Process for converting gaseous alkanes to olefins and liquid hydrocarbons

ABSTRACT

A process for converting gaseous alkanes to olefins and liquid hydrocarbons wherein a gaseous feed containing alkanes is reacted with a dry bromine vapor to form alkyl bromides and hydrobromic acid vapor. The mixture of alkyl bromides and hydrobromic acid are then reacted over a synthetic crystalline alumino-silicate catalyst, such as an X or Y type zeolite, at a temperature of from about 250° C. to about 500° C. so as to form olefins, higher molecular weight hydrocarbons and hydrobromic acid vapor. Various methods are disclosed to remove the hydrobromic acid vapor from the higher molecular weight hydrocarbons and to generate bromine from the hydrobromic acid for use in the process.

REFERENCE TO RELATED PATENT APPLICATION:

This application is a continuation-in-part of copending U.S. patentapplication Ser. No. 11/101,886 filed on Apr. 8, 2005 and entitled“Process for Converting Gaseous Alkanes to Liquid Hydrocarbons” which isa continuation-in-part of copending U.S. patent application Ser. No.10/826,885 filed on Apr. 16, 2004 and entitled “Process for ConvertingGaseous Alkanes to Liquid Hydrocarbons”.

BACKGROUND OF THE INVENTION

1. Field of the Invention

The present invention relates to a process for converting lowermolecular weight, gaseous alkanes to olefins that are useful as monomersand intermediaries in the production of chemicals, such as lubricant andfuel additives, and higher molecular weight hydrocarbons, and moreparticularly, to a process wherein a gas containing lower molecularweight alkanes is reacted with a dry bromine vapor to form alkylbromides and hydrobromic acid which in turn are reacted over acrystalline alumino-silicate catalyst to form olefins and highermolecular weight hydrocarbons.

2. Description of Related Art

Natural gas which is primarily composed of methane and other lightalkanes has been discovered in large quantities throughout the world.Many of the locales in which natural gas has been discovered are farfrom populated regions which have significant gas pipelineinfrastructure or market demand for natural gas. Due to the low densityof natural gas, transportation thereof in gaseous form by pipeline or ascompressed gas in vessels is expensive. Accordingly, practical andeconomic limits exist to the distance over which natural gas may betransported in gaseous form exist. Cryogenic liquefaction of natural gas(LNG) is often used to more economically transport natural gas overlarge distances. However, this LNG process is expensive and there arelimited regasification facilities in only a few countries that areequipped to import LNG.

Another use of methane found in natural gas is as feed to processes forthe production of methanol. Methanol is made commercially via conversionof methane to synthesis gas (CO and H₂) at high temperatures(approximately 1000° C.) followed by synthesis at high pressures(approximately 100 atmospheres). There are several types of technologiesfor the production of synthesis gas (CO and H₂) from methane. Amongthese are steam-methane reforming (SMR), partial oxidation (POX),autothermal reforming (ATR), gas-heated reforming (GHR), and variouscombinations thereof. SMR and GHR operate at high pressures andtemperatures, generally in excess of 600° C., and require expensivefurnaces or reactors containing special heat and corrosion-resistantalloy tubes filled with expensive reforming catalyst. POX and ATRprocesses operate at high pressures and even higher temperatures,generally in excess of 1000° C. As there are no known practical metalsor alloys that can operate at these temperatures, complex and costlyrefractory-lined reactors and high-pressure waste-heat boilers to quench& cool the synthesis gas effluent are required. Also, significantcapital cost and large amounts of power are required for compression ofoxygen or air to these high-pressure processes. Thus, due to the hightemperatures and pressures involved, synthesis gas technology isexpensive, resulting in a high cost methanol product which limitshigher-value uses thereof, such as for chemical feedstocks and solvents.Furthermore production of synthesis gas is thermodynamically andchemically inefficient, producing large excesses of waste heat andunwanted carbon dioxide, which tends to lower the conversion efficiencyof the overall process. Fischer-Tropsch Gas-to-Liquids (GTL) technologycan also be used to convert synthesis gas to heavier liquidhydrocarbons, however investment cost for this process is even higher.In each case, the production of synthesis gas represents a largefraction of the capital costs for these methane conversion processes.

Numerous alternatives to the conventional production of synthesis gas asa route to methanol or synthetic liquid hydrocarbons have been proposed.However, to date, none of these alternatives has attained commercialstatus for various reasons. Some of the previous alternative prior-artmethods, such as disclosed in U.S. Pat. Nos. 5,243,098 or 5,334,777 toMiller, teach reacting a lower alkane, such as methane, with a metallichalide to form a metalous halide and hydrohalic acid which are in turnreduced with magnesium oxide to form the corresponding alkanol. However,halogenation of methane using chlorine as the preferred halogen resultsin poor selectivity to the monomethyl halide (CH₃Cl), resulting inunwanted by-products such as CH₂Cl₂ and CHCl₃ which are difficult toconvert or require severe limitation of conversion per pass and hencevery high recycle rates.

Other prior art processes propose the catalytic chlorination orbromination of methane as an alternative to generation of synthesis gas(CO and H₂). To improve the selectivity of a methane halogenation stepin an overall process for the production of methanol, U.S. Pat. No.5,998,679 to Miller teaches the use of bromine, generated by thermaldecomposition of a metal bromide, to brominate alkanes in the presenceof excess alkanes, which results in improved selectivity tomono-halogenated intermediates such as methyl bromide. To avoid thedrawbacks of utilizing fluidized beds of moving solids, the processutilizes a circulating liquid mixture of metal chloride hydrates andmetal bromides. Processes described in U.S. Pat. No. 6,462,243 B1, U.S.Pat. No. 6,472,572 B1, and U.S. Pat. No. 6,525,230 to Grosso are alsocapable of attaining higher selectivity to mono-halogenatedintermediates by the use of bromination. The resulting alkyl bromidesintermediates such as methyl bromide, are further converted to thecorresponding alcohols and ethers, by reaction with metal oxides incirculating beds of moving solids. Another embodiment of U.S. Pat. No.6,525,230 avoids the drawbacks of moving beds by utilizing a zonedreactor vessel containing a fixed bed of metal oxide/metal bromide thatis operated cyclically in four steps. These processes also tend toproduce substantial quantities of dimethylether (DME) along with anyalcohol. While DME is a promising potential diesel engine fuelsubstitute, as of yet, there currently exists no substantial market forDME, and hence an expensive additional catalytic process conversion stepwould be required to convert DME into a currently marketable product.Other processes have been proposed which circumvent the need forproduction of synthesis gas, such as U.S. Pat. Nos. 4,655,893 and4,467,130 to Olah in which methane is catalytically condensed intogasoline-range hydrocarbons via catalytic condensation using superacidcatalysts. However, none of these earlier alternative approaches haveresulted in commercial processes.

It is known that substituted alkanes, in particular methanol, can beconverted to olefins and gasoline boiling-range hydrocarbons overvarious forms of crystalline alumino-silicates also known as zeolites.In the Methanol to Gasoline (MTG) process, a shape selective zeolitecatalyst, ZSM-5, is used to convert methanol to gasoline. Coal ormethane gas can thus be converted to methanol using conventionaltechnology and subsequently converted to gasoline. However due to thehigh cost of methanol production, and at current or projected prices forgasoline, the MTG process is not considered economically viable. Thus, aneed exists for an economic process for the for the conversion ofmethane and other alkanes found in natural gas to olefins and highermolecular weight hydrocarbons which, due to their higher density andvalue, are more economically transported thereby significantly aidingdevelopment of remote natural gas reserves. A further need exists for aprocess for converting alkanes present in natural gas which isrelatively inexpensive, safe and simple in operation.

SUMMARY OF THE INVENTION

To achieve the foregoing and other objects, and in accordance with thepurposes of the present invention, as embodied and broadly describedherein, one characterization of the present invention is a process isprovided for converting gaseous alkanes to to olefins and liquidhydrocarbons. A gaseous feed having lower molecular weight alkanes isreacted with bromine vapor to form alkyl bromides and hydrobromic acid.The alkyl bromides and hydrobromic acid are then reacted in the presenceof a synthetic crystalline alumino-silicate catalyst and at atemperature sufficient to form olefins and hydrobromic acid vapor.

In another characterization of the present invention, a process isprovided for converting gaseous lower molecular weight alkanes toolefins which comprises reacting a gaseous feed containing lowermolecular weight alkanes with bromine vapor to form alkyl bromides andhydrobromic acid. The alkyl bromides and hydrobromic acid are thenreacted in the presence of a synthetic crystalline alumino-silicatecatalyst to form olefins and hydrobromic acid. The process also includesconverting the hydrobromic acid to bromine.

In still another characterization of the present invention, a process isprovided for converting gaseous alkanes to olefins. A gaseous feedhaving lower molecular weight alkanes is reacted with bromine vapor toform alkyl bromides and hydrobromic acid. The alkyl bromides are reactedwith hydrobromic acid in the presence of a synthetic crystallinealumino-silicate catalyst and at a temperature sufficient to formolefins and hydrobromic acid vapor. Hydrobromic acid vapor is removedfrom the olefins by reacting the hydrobromic acid vapor with a metaloxide to form a metal bromide and steam.

BRIEF DESCRIPTION OF THE DRAWINGS

The accompanying drawings, which are incorporated in and form a part ofthe specification, illustrate the embodiments of the present inventionand, together with the description, serve to explain the principles ofthe invention.

In the drawings:

FIG. 1 is a simplified block flow diagram of the process of the presentinvention;

FIG. 2 is a schematic view of one embodiment of the process of thepresent invention;

FIG. 3 is a schematic view of another embodiment of process of thepresent invention;

FIG. 4A is schematic view of another embodiment of the process of thepresent invention;

FIG. 4B is a schematic view of the embodiment of the process of thepresent invention illustrated in FIG. 4A depicting an alternativeprocessing scheme which may be employed when oxygen is used in lieu ofair in the oxidation stage;

FIG. 5A is a schematic view of the embodiment of the process of thepresent invention illustrated in FIG. 4A with the flow through the metaloxide beds being reversed;

FIG. 5B is a schematic view of the embodiment of the process of thepresent invention illustrated in FIG. 5A depicting an alternativeprocessing scheme which may be employed when oxygen is used in lieu ofair in the oxidation stage;

FIG. 6A is a schematic view of another embodiment of the process of thepresent invention;

FIG. 6B is a schematic view of the embodiment of the process of thepresent invention illustrated in FIG. 6A depicting an alternativeprocessing scheme which may be employed when oxygen is used in lieu ofair in the oxidation stage;

FIG. 7 is a schematic view of another embodiment of the process of thepresent invention;

FIG. 8 is a schematic view of the embodiment of the process of thepresent invention illustrated in FIG. 7 with the flow through the metaloxide beds being reversed; and

FIG. 9 is a schematic view of another embodiment of the process of thepresent invention.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

As utilized throughout this description, the term “lower molecularweight alkanes” refers to methane, ethane, propane, butane, pentane ormixtures thereof. As also utilized throughout this description, “alkylbromides” refers to mono, di, and tri brominated alkanes. Also, the feedgas in lines 11 and 111 in the embodiments of the process of the presentinvention as illustrated in FIGS. 2 and 3, respectively, is preferablynatural gas which may be treated to remove sulfur compounds and carbondioxide. In any event, it is important to note that small amounts ofcarbon dioxide, e.g. less than about 2 mol %, can be tolerated in thefeed gas to the process of the present invention.

A block flow diagram generally depicting the process of the presentinvention is illustrated in FIG. 1, while specific embodiments of theprocess of the present invention are illustrated in FIGS. 2 and 3.Referring to FIG. 2, a gas stream containing lower molecular weightalkanes, comprised of a mixture of a feed gas plus a recycled gas streamat a pressure in the range of about 1 bar to about 30 bar, istransported or conveyed via line, pipe or conduit 62, mixed with drybromine liquid transported via line 25 and pump 24, and passed to heatexchanger 26 wherein the liquid bromine is vaporized. The mixture oflower molecular weight alkanes and dry bromine vapor is fed to reactor30. Preferably, the molar ratio of lower molecular weight alkanes to drybromine vapor in the mixture introduced into reactor 30 is in excess of2.5:1. Reactor 30 has an inlet pre-heater zone 28 which heats themixture to a reaction initiation temperature in the range of about 250°C. to about 400° C.

In first reactor 30, the lower molecular weight alkanes are reactedexothermically with dry bromine vapor at a relatively low temperature inthe range of about 250° C. to about 600° C., and at a pressure in therange of about 1 bar to about 30 bar to produce gaseous alkyl bromidesand hydrobromic acid vapors. The upper limit of the operatingtemperature range is greater than the upper limit of the reactioninitiation temperature range to which the feed mixture is heated due tothe exothermic nature of the bromination reaction. In the case ofmethane, the formation of methyl bromide occurs in accordance with thefollowing general reaction:CH₄ (g)+Br₂ (g)→CH₃Br (g)+HBr (g)

This reaction occurs with a significantly high degree of selectivity tomethyl bromide. For example, in the case of bromination of methane, amethane to bromine ratio of about 4.5:1 increases the selectivity to themono-halogenated methyl bromide. Small amounts of dibromomethane andtribromomethane are also formed in the bromination reaction. Higheralkanes, such as ethane, propane and butane, are also readily brominatedresulting in mono and multiple brominated species such as ethylbromides, propyl bromides and butyl bromides. If an alkane to bromineratio of significantly less than about 2.5 to 1 is utilized, a lowerselectivity to methyl bromide occurs and significant formation ofundesirable carbon soot is observed. Further, the dry bromine vapor thatis feed into first reactor 30 is substantially water-free. Applicant hasdiscovered that elimination of substantially all water vapor from thebromination step in first reactor 30 substantially eliminates theformation of unwanted carbon dioxide thereby increasing the selectivityof alkane bromination to alkyl bromides and eliminating the large amountof waste heat generated in the formation of carbon dioxide from alkanes.

The effluent that contains alkyl bromides and hydrobromic acid iswithdrawn from the first reactor via line 31 and is partially cooled toa temperature in the range of about 150° C. to about 450° C. in heatexchanger 32 before flowing to a second reactor 34. In second reactor34, the alkyl bromides are reacted exothermically at a temperature rangeof from about 250° C. to about 500° C., and a pressure in the range ofabout 1 to 20 bar, over a fixed bed 33 of crystalline alumino-silicatecatalyst, preferably a zeolite catalyst, and most preferably an X typeor Y type zeolite catalyst. A preferred zeolite is 10 X or Y typezeolite, although other zeolites with differing pore sizes andacidities, which are synthesized by varying the alumina-to-silica ratiomay be used in the process of the present invention as will be evidentto a skilled artisan. Although the zeolite catalyst is preferably usedin a protonic form, a sodium form or a mixed protonic/sodium form, thezeolite may also be modified by ion exchange with other alkali metalcations, such as Li, K or Cs, with alkali-earth metal cations, such asMg, Ca, Sr or Ba, or with transition metal cations, such as Ni, Mn, V,W, or to the hydrogen form. These various alternative cations have aneffect of shifting reaction selectivity. Other zeolite catalysts havingvarying pore sizes and acidities, which are synthesized by varying thealumina-to-silica ratio may be used in the second reactor 34 as will beevident to a skilled artisan. In this reactor, the alkyl bromides arereacted to produce a mixture of olefins and various higher molecularweight hydrocarbons, and additional hydrobromic acid vapor.

The temperature at which the second reactor 34 is operated is animportant parameter in determining the selectivity of the reaction toolefins and various higher molecular weight hydrocarbons. It ispreferred to operated second reactor 34 at a temperature within therange of about 250° to 500° C. Temperatures above about 450° C. in thesecond reactor result in increased yields of light hydrocarbons, such asundesirable methane and also deposition of coke, whereas lowertemperatures increase yields of ethylene, propylene, butylene andheavier molecular weight hydrocarbon products. Also it is noted thatmethyl bromide appears to be more reactive over a lower temperaturerange relative to methyl chloride or other substituted methyl compoundssuch as methanol. Notably, in the case of the alkyl bromide reactionover the preferred 10 X zeolite catalyst, cyclization reactions alsooccur such that the C7+ fractions contain substantial substitutedaromatics. At increasing temperatures approaching 400° C., methylbromide conversion increases towards 90% or greater, however selectivitytowards C₅+ products decreases and selectivity towards lighter products,particularly olefins increases. At temperatures exceeding 550° C., ahigh conversion of methyl bromide to methane and carbonaceous, cokeoccurs. In the preferred operating temperature range of between about300° C. and 450° C., as a byproduct of the reaction, a lesser amount ofcoke will build up on the catalyst over time during operation, causing adecline in catalyst activity over a range of hours, up to hundreds ofhours, depending on the reaction conditions and the composition of thefeed gas. It is believed that higher reaction temperatures above about400° C., associated with the formation of methane favor the thermalcracking of alkyl bromides and formation of carbon or coke and hence anincrease in the rate of deactivation of the catalyst. Conversely,temperatures at the lower end of the range, particularly below about300° C. may also contribute to coking due to a reduced rate ofdesorption of heavier products from the catalyst. Hence, operatingtemperatures within the range of about 250° C. to about 500° C., butpreferably in the range of about 300° C. to about 450° C. in the secondreactor 34 balance increased selectivity of the desired olefins and C₅+products and lower rates of deactivation due to carbon formation,against higher conversion per pass, which minimizes the quantity ofcatalyst, recycle rates and equipment size required.

The catalyst may be periodically regenerated in situ, by isolatingreactor 34 from the normal process flow, purging with an inert gas vialine 70 at a pressure in a range from about 1 to about 5 bar at anelevated temperature in the range of about 400° C. to about 650° C. toremove unreacted material adsorbed on the catalyst insofar as ispractical, and then subsequently oxidizing the deposited carbon to CO₂by addition of air or inert gas-diluted oxygen to reactor 34 via line 70at a pressure in the range of about 1 bar to about 5 bar at an elevatedtemperature in the range of about 400° C. to about 650° C. Carbondioxide and residual air or inert gas is vented from reactor 34 via line75 during the regeneration period.

The effluent which comprises olefins, the higher molecular weighthydrocarbons and hydrobromic acid is withdrawn from the second reactor34 via line 35 and is cooled to a temperature in the range of 0° C. toabout 100° C. in exchanger 36 and combined with vapor effluent in line12 from hydrocarbon stripper 47, which contains feed gas and residualhydrocarbons stripped-out by contact with the feed gas in hydrocarbonstripper 47. The combined vapor mixture is passed to a scrubber 38 andcontacted with a concentrated aqueous partially-oxidized metal bromidesalt solution containing metal hydroxide and/or metal oxide and/or metaloxy-bromide species, which is transported to scrubber 38 via line 41.The preferred metal of the bromide salt is Fe(III), Cu(II) or Zn(II), ormixtures thereof, as these are less expensive and readily oxidize atlower temperatures in the range of about 120° C. to about 180° C.,allowing the use of glass-lined or fluorpolymer-lined equipment;although Co(II), Ni(II), Mn(II), V(II), Cr(II) or othertransition-metals which form oxidizable bromide salts may be used in theprocess of the present invention. Alternatively, alkaline-earth metalswhich also form oxidizable bromide salts, such as Ca (II) or Mg(II) maybe used. Any liquid hydrocarbons condensed in scrubber 38 may be skimmedand withdrawn in line 37 and added to liquid hydrocarbons exiting theproduct recovery unit 52 in line 54. Hydrobromic acid is dissolved inthe aqueous solution and neutralized by the metal hydroxide and/or metaloxide and/or metal oxy-bromide species to yield metal bromide salt insolution and water which is removed from the scrubber 38 via line 44.

The residual vapor phase containing olefins and the higher molecularweight hydrocarbons that is removed as effluent from the scrubber 38 isforwarded via line 39 to dehydrator 50 to remove substantially all watervia line 53 from the vapor stream. The water is then removed from thedehydrator 50 via line 53. The dried vapor stream containing olefins andthe higher molecular weight hydrocarbons is further passed via line 51to product recovery unit 52 to recover olefins and the C₅+gasoline-range hydrocarbon fraction as a liquid product in line 54. Anyconventional method of dehydration and liquids recovery, such assolid-bed dessicant adsorption followed by refrigerated condensation,cryogenic expansion, or circulating absorption oil or other solvent, asused to process natural gas or refinery gas streams, and to recoverolefinic hydrocarbons, as will be evident to a skilled artisan, may beemployed in the process of the present invention. The residual vaporeffluent from product recovery unit 52 is then split into a purge stream57 which may be utilized as fuel for the process and a recycled residualvapor which is compressed via compressor 58. The recycled residual vapordischarged from compressor 58 is split into two fractions. A firstfraction that is equal to at least 2.5 times the feed gas molar volumeis transported via line 62 and is combined with dry liquid bromineconveyed by pump 24, heated in exchanger 26 to vaporize the bromine andfed into first reactor 30. The second fraction is drawn off of line 62via line 63 and is regulated by control valve 60, at a rate sufficientto dilute the alkyl bromide concentration to reactor 34 and absorb theheat of reaction such that reactor 34 is maintained at the selectedoperating temperature, preferably in the range of about 300° C. to about450° C. in order to maximize conversion versus selectivity and tominimize the rate of catalyst deactivation due to the deposition ofcarbon. Thus, the dilution provided by the recycled vapor effluentpermits selectivity of bromination in the first reactor 30 to becontrolled in addition to moderating the temperature in second reactor34.

Water containing metal bromide salt in solution which is removed fromscrubber 38 via line 44 is passed to hydrocarbon stripper 47 whereinresidual dissolved hydrocarbons are stripped from the aqueous phase bycontact with incoming feed gas transported via line 11. The strippedaqueous solution is transported from hydrocarbon stripper 47 via line 65and is cooled to a temperature in the range of about 0° C. to about 70°C. in heat exchanger 46 and then passed to absorber 48 in which residualbromine is recovered from vent stream in line 67. The aqueous solutioneffluent from scrubber 48 is transported via line 49 to a heat exchanger40 to be preheated to a temperature in the range of about 100° C. toabout 600° C., and most preferably in the range of about 120° C. toabout 180° C. and passed to third reactor 16. Oxygen or air is deliveredvia line 10 by blower or compressor 13 at a pressure in the range ofabout ambient to about 5 bar to bromine stripper 14 to strip residualbromine from water which is removed from stripper 14 in line 64 and iscombined with water stream 53 from dehydrator 50 to form water effluentstream in line 56 which is removed from the process. The oxygen or airleaving bromine stripper 14 is fed via line 15 to reactor 16 whichoperates at a pressure in the range of about ambient to about 5 bar andat a temperature in the range of about 100° C. to about 600° C., butmost preferably in the range of about 120° C. to about 180° C. so as tooxidize an aqueous metal bromide salt solution to yield elementalbromine and metal hydroxide and/or metal oxide and or metal oxy-bromidespecies. As stated above, although Co(II), Ni(II), Mn(II), V(II), Cr(II)or other transition-metals which form oxidizable bromide salts can beused, the preferred metal of the bromide salt is Fe(III), Cu(II), orZn(II), or mixtures thereof, as these are less expensive and readilyoxidize at lower temperatures in the range of about 120° C. to about180° C., allowing the use of glass-lined or fluorpolymer-linedequipment. Alternatively, alkaline-earth metals which also formoxidizable bromide salts, such as Ca (II) or Mg(II) could be used.

Hydrobromic acid reacts with the metal hydroxide and/or metal oxideand/or metal oxy-bromide species so formed to once again yield the metalbromide salt and water. Heat exchanger 18 in reactor 16 supplies heat tovaporize water and bromine. Thus, the overall reactions result in thenet oxidation of hydrobromic acid produced in first reactor 30 andsecond reactor 34 to elemental bromine and steam in the liquid phasecatalyzed by the metal bromide/metal oxide or metal hydroxide operatingin a catalytic cycle. In the case of the metal bromide being Fe(III)Br3,the reactions are believed to be:Fe(+3a)+6Br(−a)+3H(+a)+3/2O2(g)=3Br2(g)+Fe(OH)3   1)3HBr(g)+H₂O=3H(+a)+3Br(−a)+H₂O   2)3H(+a)+3Br(−a)+Fe(OH)3=Fe(+3a)+3Br(−a)+3H₂O   3)

The elemental bromine and water and any residual oxygen or nitrogen (ifair is utilized as the oxidant) leaving as vapor from the outlet ofthird reactor 16 via line 19, are cooled in condenser 20 at atemperature in the range of about 0° C. to about 70° C. and a pressurein the range of about ambient to 5 bar to condense the bromine and waterand passed to three-phase separator 22. In three-phase separator 22,since liquid water has a limited solubility for bromine, on the order ofabout 3% by weight, any additional bromine which is condensed forms aseparate, denser liquid bromine phase. The liquid bromine phase,however, has a notably lower solubility for water, on the order of lessthan 0.1%. Thus a substantially dry bromine vapor can be easily obtainedby condensing liquid bromine and water, decanting water by simplephysical separation and subsequently re-vaporizing liquid bromine.

Liquid bromine is pumped in line 25 from three-phase separator 22 viapump 24 to a pressure sufficient to mix with vapor stream 62. Thusbromine is recovered and recycled within the process. The residualoxygen or nitrogen and any residual bromine vapor which is not condensedexits three-phase separator 22 and is passed via line 23 to brominescrubber 48, wherein residual bromine is recovered by solution into andby reaction with reduced metal bromides in the aqueous metal bromidesolution stream 65. Water is removed from separator 22 via line 27 andintroduced into stripper 14.

In another embodiment of the invention, referring to FIG. 3, a gasstream containing lower molecular weight alkanes, comprised of mixtureof a feed gas plus a recycled gas stream at a pressure in the range ofabout 1 bar to about 30 bar, is transported or conveyed via line, pipeor conduit 162, mixed with dry bromine liquid transported via pump 124and passed to heat exchanger 126 wherein the liquid bromine isvaporized. The mixture of lower molecular weight alkanes and dry brominevapor is fed to reactor 130. Preferably, the molar ratio of lowermolecular weight alkanes to dry bromine vapor in the mixture introducedinto reactor 130 is in excess of 2.5:1 Reactor 130 has an inletpre-heater zone 128 which heats the mixture to a reaction initiationtemperature in the range of about 250° C. to about 400° C. In firstreactor 130, the lower molecular weight alkanes are reactedexothermically with dry bromine vapor at a relatively low temperature inthe range of about 250° C. to about 600° C., and at a pressure in therange of about 1 bar to about 30 bar to produce gaseous alkyl bromidesand hydrobromic acid vapors. The upper limit of the operatingtemperature range is greater than the upper limit of the reactioninitiation temperature range to which the feed mixture is heated due tothe exothermic nature of the bromination reaction. In the case ofmethane, the formation of methyl bromide occurs in accordance with thefollowing general reaction:CH₄ (g)+Br₂ (g)→CH₃Br (g)+HBr (g)This reaction occurs with a significantly high degree of selectivity tomethyl bromide. For example, in the case of bromine reacting with amolar excess of methane at a methane to bromine ratio of 4.5:1, a highselectivity to the mono-halogenated methyl bromide occurs. Small amountsof dibromomethane and tribromomethane are also formed in the brominationreaction. Higher alkanes, such as ethane, propane and butane, are alsoreadily brominated resulting in mono and multiple brominated speciessuch as ethyl bromides, propyl bromides and butyl bromides. If an alkaneto bromine ratio of significantly less than 2.5 to 1 is utilized,substantially lower selectivity to methyl bromide substantially occursand significant formation of undesirable carbon soot is observed.Further, the dry bromine vapor that is feed into first reactor 130 issubstantially water-free. Applicant has discovered that elimination ofsubstantially all water vapor from the bromination step in first reactor130 substantially eliminates the formation of unwanted carbon dioxidethereby increasing the selectivity of alkane bromination to alkylbromides and eliminating the large amount of waste heat generated in theformation of carbon dioxide from alkanes.

The effluent that contains alkyl bromides and hydrobromic acid iswithdrawn from the first reactor 130 via line 131 and is partiallycooled to a temperature in the range of about 150° C. to 450° C. in heatexchanger 132 before flowing to a second reactor 134. In second reactor134, the alkyl bromides are reacted exothermically at a temperaturerange of from about 250° C. to about 500° C., and a pressure in therange of about 1 bar to 30 bar, over a fixed bed of crystallinealumino-silicate catalyst, preferably a zeolite catalyst, and mostpreferably an X type or Y type zeolite catalyst. A preferred zeolite is10 X or Y type zeolite, although other zeolites with differing poresizes and acidities, which are synthesized by varying thealumina-to-silica ratio may be used in the process of the presentinvention as will be evident to a skilled artisan. Although the zeolitecatalyst is preferably used in a protonic form, a sodium form or a mixedprotonic/sodium form, such as Li, K or Cs, with alkali-earth metalcations, such as Mg, Ca, Sr or Ba, or with transition metal cations,such as Ni, Mn, V, W, or to the hydrogen form. These various alternativecations have an effect of shifting reaction selectivity. Other zeolitecatalysts having varying pore sizes and acidities, which are synthesizedby varying the alumina-to-silica ratio may be used in the second reactor134 as will be evident to a skilled artisan. In this reactor, the alkylbromides are reacted to produce a mixture of olefins and highermolecular weight hydrocarbons and additional hydrobromic acid vapor.

The temperature at which the second reactor 134 is operated is animportant parameter in determining the selectivity of the reaction toolefins and various higher molecular weight liquid hydrocarbons. It ispreferred to operate second reactor 134 at a temperature within therange of about 250° C. to 500° C. , but more preferably within the rangeof about 300° C. to 450° C. Temperatures above about 450° C. in thesecond reactor result in increased yields of light hydrocarbons, such asundesirable methane and carbonaceous coke, whereas lower temperaturesincrease yields of olefins such as ethylene, propylene and butylene andheavier molecular weight hydrocarbon products. Notably, in the case ofalkyl bromides reacting over the preferred 10 X zeolite catalyst,cyclization reactions occur such that the C₇+ fractions produced containsubstantial substituted aromatics. At increasing temperaturesapproaching 400° C., methyl bromide conversion increases towards 90% orgreater, however selectivity towards C₅+ products decreases andselectivity towards lighter products, particularly olefins, increases.At temperatures exceeding 550° C. almost complete conversion of methylbromide to methane and coke occurs. In the preferred range of operatingtemperatures of about 300° C. to 450° C., as a byproduct of thereaction, a small amount of carbon will build up on the catalyst overtime during operation, causing a decline in catalyst activity over arange of several hundred hours, depending on the reaction conditions andfeed gas composition. It is observed that higher reaction temperaturesabove about 400° C. favor the thermal cracking of alkyl bromides withformation of carbon and hence increases the rate of deactivation of thecatalyst. Conversely, operation at the lower end of the temperaturerange, particularly below about 300° C. may also promote coking, likelyto the reduced rate of desorption of hydrocarbon products. Hence,operating temperatures within the range of about 250° C. to 500° C. butmore preferably in the range of about 300° C. to 450° C. in the secondreactor 134 balance increased selectivity towards the desired olefin andC₅+ products and lower rates of deactivation due to carbon formation,against higher conversion per pass, which minimizes the quantity ofcatalyst, recycle rates and equipment size required.

The catalyst may be periodically regenerated in situ, by isolatingreactor 134 from the normal process flow, purging with an inert gas vialine 170 at a pressure in the range of about 1 bar to about 5 bar and anelevated temperature in the range of 400° C. to 650° C. to removeunreacted material adsorbed on the catalyst insofar as is practical, andthen subsequently oxidizing the deposited carbon to CO₂ by addition ofair or inert gas-diluted oxygen via line 170 to reactor 134 at apressure in the range of about 1 bar to about 5 bar and an elevatedtemperature in the range of 400° C. to 650° C. Carbon dioxide andresidual air or inert gas are vented from reactor 134 via line 175during the regeneration period.

The effluent which comprises olefins, the higher molecular weighthydrocarbons and hydrobromic acid is withdrawn from the second reactor134 via line 135, cooled to a temperature in the range of about 0° C. toabout 100° C. in exchanger 136, and combined with vapor effluent in line112 from hydrocarbon stripper 147. The mixture is then passed to ascrubber 138 and contacted with a stripped, recirculated water that istransported to scrubber 138 in line 164 by any suitable means, such aspump 143, and is cooled to a temperature in the range of about 0° C. toabout 50° C. in heat exchanger 155. Any liquid hydrocarbon productcondensed in scrubber 138 may be skimmed and withdrawn as stream 137 andadded to liquid hydrocarbon product 154. Hydrobromic acid is dissolvedin scrubber 138 in the aqueous solution which is removed from thescrubber 138 via line 144, and passed to hydrocarbon stripper 147wherein residual hydrocarbons dissolved in the aqueous solution arestripped-out by contact with feed gas 111. The stripped aqueous phaseeffluent from hydrocarbon stripper 147 is cooled to a temperature in therange of about 0° C. to about 50° C. in heat exchanger 146 and thenpassed via line 165 to absorber 148 in which residual bromine isrecovered from vent stream 167.

The residual vapor phase containing olefins and the higher molecularweight hydrocarbon is removed as effluent from the scrubber 138 andforwarded to dehydrator 150 to remove substantially all water from thegas stream. The water is then removed from the dehydrator 150 via line153. The dried gas stream containing olefins and the higher molecularweight hydrocarbons is further passed via line 151 to product recoveryunit 152 to recover olefins and the C₅+ gasoline range hydrocarbonfraction as a liquid product in line 154. Any conventional method ofdehydration and liquids recovery such as solid-bed dessicant adsorptionfollowed by, for example, refrigerated condensation, cryogenicexpansion, or circulating absorption oil, or other solvents as used toprocess natural gas or refinery gas streams and recover olefinichydrocarbons, as known to a skilled artisan, may be employed in theimplementation of this invention. The residual vapor effluent fromproduct recovery unit 152 is then split into a purge stream 157 that maybe utilized as fuel for the process and a recycled residual vapor whichis compressed via compressor 158. The recycled residual vapor dischargedfrom compressor 158 is split into two fractions. A first fraction thatis equal to at least 2.5 times the feed gas volume is transported vialine 162, combined with the liquid bromine conveyed in line 125 andpassed to heat exchanger 126 wherein the liquid bromine is vaporized andfed into first reactor 130. The second fraction which is drawn off line162 via line 163 and is regulated by control valve 160, at a ratesufficient to dilute the alkyl bromide concentration to reactor 134 andabsorb the heat of reaction such that reactor 134 is maintained at theselected operating temperature, preferably in the range of about 300° C.to about 450° C. in order to maximize conversion vs. selectivity and tominimize the rate of catalyst deactivation due to the deposition ofcarbon. Thus, the dilution provided by the recycled vapor effluentpermits selectivity of bromination in the first reactor 130 to becontrolled in addition to moderating the temperature in second reactor134.

Oxygen, oxygen enriched air or air 110 is delivered via blower orcompressor 113 at a pressure in the range of about ambient to about 5bar to bromine stripper 114 to strip residual bromine from water whichleaves stripper 114 via line 164 and is divided into two portions. Thefirst portion of the stripped water is recycled via line 164, cooled inheat exchanger 155 to a temperature in the range of about 20° C. toabout 50° C., and maintained at a pressure sufficient to enter scrubber138 by any suitable means, such as pump 143. The portion of water thatis recycled is selected such that the hydrobromic acid solution effluentremoved from scrubber 138 via line 144 has a concentration in the rangefrom about 10% to about 50% by weight hydrobromic acid, but morepreferably in the range of about 30% to about 48% by weight to minimizethe amount of water which must be vaporized in exchanger 141 andpreheater 119 and to minimize the vapor pressure of HBr over theresulting acid. A second portion of water from stripper 114 is removedfrom line 164 and the process via line 156.

The dissolved hydrobromic acid that is contained in the aqueous solutioneffluent from scrubber 148 is transported via line 149 and is combinedwith the oxygen, oxygen enriched air or air leaving bromine stripper 114in line 115. The combined aqueous solution effluent and oxygen, oxygenenriched air or air is passed to a first side of heat exchanger 141 andthrough preheater 119 wherein the mixture is preheated to a temperaturein the range of about 100° C. to about 600° C. and most preferably inthe range of about 120° C. to about 180° C. and passed to third reactor117 that contains a metal bromide salt or metal oxide. The preferredmetal of the bromide salt or metal oxide is Fe(III), Cu(II) or Zn(II)although Co(II), Ni(II), Mn(II), V(II), Cr(II) or othertransition-metals which form oxidizable bromide salts can be used.Alternatively, alkaline-earth metals which also form oxidizable bromidesalts, such as Ca (II) or Mg(II) could be used. The metal bromide saltin the oxidation reactor 117 can be utilized as a concentrated aqueoussolution or preferably, the concentrated aqueous salt solution may beimbibed into a porous, high surface area, acid resistant inert supportsuch as a silica gel. More preferably, the oxide form of the metal in arange of 10 to 20% by weight is deposited on an inert support such asalumina with a specific surface area in the range of 50 to 200 m2/g. Theoxidation reactor 117 operates at a pressure in the range of aboutambient to about 5 bar and at a temperature in the range of about 100°C. to 600° C., but most preferably in the range of about 120° C. to 180°C.; therein, the metal bromide is oxidized by oxygen, yielding elementalbromine and metal hydroxide, metal oxide or metal oxy-bromide speciesor, metal oxides in the case of the supported metal bromide salt ormetal oxide operated at higher temperatures and lower pressures at whichwater may primarily exist as a vapor. In either case, the hydrobromicacid reacts with the metal hydroxide, metal oxy-bromide or metal oxidespecies and is neutralized, restoring the metal bromide salt andyielding water. Thus, the overall reaction results in the net oxidationof hydrobromic acid produced in first reactor 130 and second reactor 134to elemental bromine and steam, catalyzed by the metal bromide/metalhydroxide or metal oxide operating in a catalytic cycle. In the case ofthe metal bromide being Fe(III)Br2 in an aqueous solution and operatedin a pressure and temperature range in which water may exist as a liquidthe reactions are believed to be:Fe(+3a)+6Br(−a)+3H(+a)+3/2O₂(g)=3Br₂(g)+Fe(OH)3   1)3HBr(g)+H₂O=3H(+a)+3Br(−a)+H₂O   2)3H(+a)+3Br(−a)+Fe(OH)3=Fe(+3a)+3Br(−a)+3H₂O   3)In the case of the metal bromide being Cu(II)Br2 supported on an inertsupport and operated at higher temperature and lower pressure conditionsat which water primarily exists as a vapor, the reactions are believedto be:2Cu(II)Br2=2Cu(I)Br+Br2(g)   1)2Cu(I)Br+O₂(g)=Br2(g)+2Cu(II)O   2)2HBr(g)+Cu(II)O=Cu(II)Br₂+H₂O(g)   3)

The elemental bromine and water and any residual oxygen or nitrogen (ifair or oxygen enriched air is utilized as the oxidant) leaving as vaporfrom the outlet of third reactor 117, are cooled in the second side ofexchanger 141 and condenser 120 to a temperature in the range of about0° C. to about 70° C. wherein the bromine and water are condensed andpassed to three-phase separator 122. In three-phase separator 122, sinceliquid water has a limited solubility for bromine, on the order of about3% by weight, any additional bromine which is condensed forms aseparate, denser liquid bromine phase. The liquid bromine phase,however, has a notably lower solubility for water, on the order of lessthan 0.1%. Thus, a substantially dry bromine vapor can be easilyobtained by condensing liquid bromine and water, decanting water bysimple physical separation and subsequently re-vaporizing liquidbromine. It is important to operate at conditions that result in thenear complete reaction of HBr so as to avoid significant residual HBr inthe condensed liquid bromine and water, as HBr increases the miscibilityof bromine in the aqueous phase, and at sufficiently highconcentrations, results in a single ternary liquid phase.

Liquid bromine is pumped from three-phase separator 122 via pump 124 toa pressure sufficient to mix with vapor stream 162. Thus the bromine isrecovered and recycled within the process. The residual air, oxygenenriched air or oxygen and any bromine vapor which is not condensedexits three-phase separator 122 and is passed via line 123 to brominescrubber 148, wherein residual bromine is recovered by dissolution intohydrobromic acid solution stream conveyed to scrubber 148 via line 165.Water is removed from the three-phase separator 122 via line 129 andpassed to stripper 114.

Thus, in accordance with all embodiments of the present invention setforth above, the metal bromide/metal hydroxide, metal oxy-bromide ormetal oxide operates in a catalytic cycle allowing bromine to be easilyrecycled within the process. The metal bromide is readily oxidized byoxygen, oxygen enriched air or air either in the aqueous phase or thevapor phase at temperatures in the range of about 100° C. to about 600°C. and most preferably in the range of about 120° C. to about 180° C. toyield elemental bromine vapor and metal hydroxide, metal oxy-bromide ormetal oxide. Operation at temperatures below about 180° C. isadvantageous, thereby allowing the use of low-cost corrosion-resistantfluoropolymer-lined equipment. Hydrobromic acid is neutralized byreaction with the metal hydroxide or metal oxide yielding steam and themetal bromide.

The elemental bromine vapor and steam are condensed and easily separatedin the liquid phase by simple physical separation, yieldingsubstantially dry bromine. The absence of significant water allowsselective bromination of alkanes, without production of CO₂ and thesubsequent efficient and selective reactions of alkyl bromides toprimarily C₂ to C₄ olefins and heavier products, the C₅+ fraction ofwhich contains substantial branched alkanes and substituted aromatics.Byproduct hydrobromic acid vapor from the bromination reaction andsubsequent reaction in reactor 134 are readily dissolved into an aqueousphase and neutralized by the metal hydroxide or metal oxide speciesresulting from oxidation of the metal bromide.

In accordance with another embodiment of the process of the presentinvention illustrated in FIG. 4A, the alkyl bromination and alkylbromide conversion stages are operated in a substantially similar mannerto those corresponding stages described with respect to FIGS. 2 and 3above. More particularly, a gas stream containing lower molecular weightalkanes, comprised of mixture of a feed gas and a recycled gas stream ata pressure in the range of about 1 bar to about 30 bar, is transportedor conveyed via line, pipe or conduits 262 and 211, respectively, andmixed with dry bromine liquid in line 225. The resultant mixture istransported via pump 224 and passed to heat exchanger 226 wherein theliquid bromine is vaporized. The mixture of lower molecular weightalkanes and dry bromine vapor is fed to reactor 230. Preferably, themolar ratio of lower molecular weight alkanes to dry bromine vapor inthe mixture introduced into reactor 230 is in excess of 2.5:1. Reactor230 has an inlet pre-heater zone 228 which heats the mixture to areaction initiation temperature in the range of 250° C. to 400° C. Infirst reactor 230, the lower molecular weight alkanes are reactedexothermically with dry bromine vapor at a relatively low temperature inthe range of about 250° C. to about 600° C., and at a pressure in therange of about 1 bar to about 30 bar to produce gaseous alkyl bromidesand hydrobromic acid vapors. The upper limit of the operatingtemperature range is greater than the upper limit of the reactioninitiation temperature range to which the feed mixture is heated due tothe exothermic nature of the bromination reaction. In the case ofmethane, the formation of methyl bromide occurs in accordance with thefollowing general reaction:CH₄ (g)+Br₂ (g)→CH₃Br (g)+HBr (g)This reaction occurs with a significantly high degree of selectivity tomethyl bromide. For example, in the case of bromine reacting with amolar excess of methane at a methane to bromine ratio of 4.5:1, a highselectivity to the mono-halogenated methyl bromide occurs. Small amountsof dibromomethane and tribromomethane are also formed in the brominationreaction. Higher alkanes, such as ethane, propane and butane, are alsoreadily bromoninated resulting in mono and multiple brominated speciessuch as ethyl bromides, propyl bromides and butyl bromides. If an alkaneto bromine ratio of significantly less than 2.5 to 1 is utilized,substantially lower selectivity to methyl bromide occurs and significantformation of undesirable carbon soot is observed. Further, the drybromine vapor that is feed into first reactor 230 is substantiallywater-free. Applicant has discovered that elimination of substantiallyall water vapor from the bromination step in first reactor 230substantially eliminates the formation of unwanted carbon dioxidethereby increasing the selectivity of alkane bromination to alkylbromides and eliminating the large amount of waste heat generated in theformation of carbon dioxide from alkanes.

The effluent that contains alkyl bromides and hydrobromic acid iswithdrawn from the first reactor 230 via line 231 and is partiallycooled to a temperature in the range of about 150° C. to 450° C. in heatexchanger 232 before flowing to a second reactor 234. In second reactor234, the alkyl bromides are reacted exothermically at a temperaturerange of from about 250° C. to about 500° C., and a pressure in therange of about 1 bar to 30 bar, over a fixed bed of crystallinealumino-silicate catalyst, preferably a zeolite catalyst, and mostpreferably, an X type or Y type zeolite catalyst. A preferred zeolite is10 X or Y type zeolite, although other zeolites with differing poresizes and acidities, which are synthesized by varying thealumina-to-silica ratio may be used in the process of the presentinvention as will be evident to a skilled artisan. Although the zeolitecatalyst is preferably used in a protonic form, a sodium form or a mixedprotonic/sodium form, the zeolite may also be modified by ion exchangewith other alkali metal cations, such as Li, K or Cs, with alkali-earthmetal cations, such as Mg, Ca, Sr or Ba, with transition metal cations,such as Ni, Mn, V, W, or to the hydrogen form. These various alternativecations have an effect of shifting reaction selectivity. Other zeolitecatalysts having varying pore sizes and acidities, which are synthesizedby varying the alumina-to-silica ratio may be used in the second reactor234 as will be evident to a skilled artisan. In this reactor, the alkylbromides are reacted to produce a mixture of higher molecular weighthydrocarbon products and additional hydrobromic acid vapor.

The temperature at which the second reactor 234 is operated is animportant parameter in determining the selectivity of the reaction toolefins and various higher molecular weight liquid hydrocarbons. It ispreferred to operate second reactor 234 at a temperature within therange of about 250° C. to about 500° C., but more preferably within therange of about 300° C. to about 450° C. Temperatures above about 450° C.in the second reactor result in increased yields of light hydrocarbons,such as undesirable methane and carbonaceous coke, whereas lowertemperatures increase yields of olefins and heavier molecular weighthydrocarbon products. Notably, in the case of alkyl bromides reactingover the preferred 10 X zeolite catalyst, cyclization reactions occursuch that the C₇+ fractions produced contain substantial substitutedaromatics. At increasing temperatures approaching 400° C., methylbromide conversion increases towards 90% or greater, however selectivitytowards C₅+ products decreases and selectivity towards lighter products,particularly olefins, increases. At temperatures exceeding 550° C.almost complete conversion of methyl bromide to methane and coke occurs.In the preferred range of operating temperatures of about 300° C. to450° C., as a byproduct of the reaction, a small amount of carbon willbuild up on the catalyst over time during operation, causing a declinein catalyst activity over a range of several hundred hours, depending onthe reaction conditions and feed gas composition. It is observed thathigher reaction temperatures above about 400° C. favor the thermalcracking of alkyl bromides with formation of carbon and hence increasesthe rate of deactivation of the catalyst. Conversely, operation at thelower end of the temperature range, particularly below about 300° C. mayalso promote coking, likely to the reduced rate of desorption ofhydrocarbon products. Hence, operating temperatures within the range ofabout 250° C. to 500° C. but more preferably in the range of about 300°C. to 450° C. in the second reactor 234 balance increased selectivitytowards the desired olefin and C₅+ products and lower rates ofdeactivation due to carbon formation, against higher conversion perpass, which minimizes the quantity of catalyst, recycle rates andequipment size required.

The catalyst may be periodically regenerated in situ, by isolatingreactor 234 from the normal process flow, purging with an inert gas vialine 270 at a pressure in the range of about 1 bar to about 5 bar and anelevated temperature in the range of about 400° C. to about 650° C. toremove unreacted material adsorbed on the catalyst insofar as ispractical, and then subsequently oxidizing the deposited carbon to CO₂by addition of air or inert gas-diluted oxygen via line 270 to reactor234 at a pressure in the range of about 1 bar to about 5 bar and anelevated temperature in the range of about 400° C. to about 650° C.Carbon dioxide and residual air or inert gas are vented from reactor 234via line 275 during the regeneration period.

The effluent which comprises olefins, the higher molecular weighthydrocarbons and hydrobromic acid is withdrawn from the second reactor234 via line 235 and cooled to a temperature in the range of about 100°C. to about 600° C. in exchanger 236. As illustrated in FIG. 4A, thecooled effluent is transported via lines 235 and 241 with valve 238 inthe opened position and valves 239 and 243 in the closed position andintroduced into a vessel or reactor 240 containing a bed 298 of a solidphase metal oxide. The metal of the metal oxide is selected formmagnesium (Mg), calcium (Ca), vanadium (V), chromium (Cr), manganese(Mn), iron (Fe), cobalt (Co), nickel (Ni), copper (Cu), zinc (Sn), ortin (Sn). The metal is selected for the impact of its physical andthermodynamic properties relative to the desired temperature ofoperation, and also for potential environmental and health impacts andcost. Preferably, magnesium, copper and iron are employed as the metal,with magnesium being the most preferred. These metals have the propertyof not only forming oxides but bromide salts as well, with the reactionsbeing reversible in a temperature range of less than about 500° C. Thesolid metal oxide is preferably immobilized on a suitableattrition-resistant support, for example a synthetic amorphous silica,such as Davicat Grade 57, manufactured by Davison Catalysts of Columbia,Md. Or more preferably, an alumina support with a specific surface areaof about 50 to 200 m2/g. In reactor 240, hydrobromic acid is reactedwith the metal oxide at temperatures below about 600° C. and preferablybetween about 100° C. to about 500° C. in accordance with the followinggeneral formula wherein M represents the metal:2HBr+MO→MBr₂+H₂OThe steam resulting from this reaction is transported together witholefins and the high molecular hydrocarbons in line 244, 218 and 216 viaopened valve 219 to heat exchanger 220 wherein the mixture is cooled toa temperature in the range of about 0° C. to about 70° C. This cooledmixture is forwarded to dehydrator 250 to remove substantially all waterfrom the gas stream. The water is then removed from the dehydrator 250via line 253. The dried gas stream containing olefins and the highermolecular weight hydrocarbons is further passed via line 251 to productrecovery unit 252 to recover olefins and the C₅+ fraction as a liquidproduct in line 254. Any conventional method of dehydration and liquidsrecovery such as solid-bed dessicant adsorption followed by, forexample, refrigerated condensation, cryogenic expansion, or circulatingabsorption oil or other solvent, as used to process natural gas orrefinery gas streams and recover olefinic hydrocarbons, as known to askilled artisan, may be employed in the implementation of thisinvention. The residual vapor effluent from product recovery unit 252 isthen split into a purge stream 257 that may be utilized as fuel for theprocess and a recycled residual vapor which is compressed via compressor258. The recycled residual vapor discharged from compressor 258 is splitinto two fractions. A first fraction that is equal to at least 1.5 timesthe feed gas volume is transported via line 262, combined with theliquid bromine and feed gas conveyed in line 225 and passed to heatexchanger 226 wherein the liquid bromine is vaporized and fed into firstreactor 230 in a manner as described above. The second fraction which isdrawn off line 262 via line 263 and is regulated by control valve 260,at a rate sufficient to dilute the alkyl bromide concentration toreactor 234 and absorb the heat of reaction such that reactor 234 ismaintained at the selected operating temperature, preferably in therange of about 300° C. to about 450° C. in order to maximize conversionvs. selectivity and to minimize the rate of catalyst deactivation due tothe deposition of carbon. Thus, the dilution provided by the recycledvapor effluent permits selectivity of bromination in the first reactor230 to be controlled in addition to moderating the temperature in secondreactor 234.

Oxygen, oxygen enriched air or air 210 is delivered via blower orcompressor 213 at a pressure in the range of about ambient to about 10bar to bromine via line 214, line 215 and valve 249 through heatexchanger 215, wherein oxygen, oxygen enriched air or air is preheatedto a temperature in the range of about 100° C. to about 500° C. to asecond vessel or reactor 246 containing a bed 299 of a solid phase metalbromide. Oxygen reacts with the metal bromide in accordance with thefollowing general reaction wherein M represents the metal:MBr₂+1/2O₂→MO+Br₂In this manner, a dry, substantially HBr free bromine vapor is producedthereby eliminating the need for subsequent separation of water orhydrobromic acid from the liquid bromine. Reactor 246 is operated below600° C., and more preferably between about 300° C. to about 500° C. Theresultant bromine vapor is transported from reactor 246 via line 247,valve 248 and line 242 to heat exchanger or condenser 221 where thebromine is condensed into a liquid. The liquid bromine is transportedvia line 242 to separator 222 wherein liquid bromine is removed via line225 and transported via line 225 to heat exchanger 226 and first reactor230 by any suitable means, such as by pump 224. The residual air orunreacted oxygen is transported from separator 222 via line 227 to abromine scrubbing unit 223, such as venturi scrubbing system containinga suitable solvent, or suitable solid adsorbant medium, as selected by askilled artisan, wherein the remaining bromine is captured. The capturedbromine is desorbed from the scrubbing solvent or adsorbant by heatingor other suitable means and the thus recovered bromine transported vialine 212 to line 225. The scrubbed air or oxygen is vented via line 229.In this manner, nitrogen and any other substantially non-reactivecomponents are removed from the system of the present invention andthereby not permitted to enter the hydrocarbon-containing portion of theprocess; also loss of bromine to the surrounding environment is avoided.

One advantage of removing the HBr by chemical reaction in accordancewith this embodiment, rather than by simple physical solubility, is thesubstantially complete scavenging of the HBr to low levels at higherprocess temperatures. Another distinct advantage is the elimination ofwater from the bromine removed thereby eliminating the need forseparation of bromine and water phases and for stripping of residualbromine from the water phase.

Reactors 240 and 246 may be operated in a cyclic fashion. As illustratedin FIG. 4A, valves 238 and 219 are operated in the open mode to permithydrobromic acid to be removed from the effluent that is withdrawn fromthe second reactor 234, while valves 248 and 249 are operated in theopen mode to permit air, oxygen enriched air or oxygen to flow throughreactor 246 to oxidize the solid metal bromide contained therein. Oncesignificant conversion of the metal oxide and metal bromide in reactors240 and 246, respectively, has occurred, these valves are closed. Atthis point, bed 299 in reactor 246 is a bed of substantially solid metalbromide, while bed 298 in reactor 240 is substantially solid metaloxide. As illustrated in FIG. 5A, valves 245 and 243 are then opened topermit oxygen, oxygen enriched air or air to flow through reactor 240 tooxidize the solid metal bromide contained therein, while valves 239 and217 are opened to permit effluent which comprises olefins, the highermolecular weight hydrocarbons and hydrobromic acid that is withdrawnfrom the second reactor 234 to be introduced into reactor 246. Thereactors are operated in this manner until significant conversion of themetal oxide and metal bromide in reactors 246 and 240, respectively, hasoccurred and then the reactors are cycled back to the flow schematicillustrated in FIG. 4A by opening and closing valves as previouslydiscussed.

When oxygen is utilized as the oxidizing gas transported in via line 210to the reactor being used to oxidize the solid metal bromide containedtherein, the embodiment of the process of the present inventionillustrated in FIGS. 4A and 5A can be modified such that the brominevapor produced from either reactor 246 (FIG. 4B) or 240 (FIG. 5B) istransported via lines 242 and 225 directly to first reactor 230. Sinceoxygen is reactive and will not build up in the system, the need tocondense the bromine vapor to a liquid to remove unreactive components,such as nitrogen, is obviated. Compressor 213 is not illustrated inFIGS. 4B and 5B as substantially all commercial sources of oxygen, suchas a commercial air separator unit, will provide oxygen to line 210 atthe required pressure. If not, a compressor 213 could be utilized toachieve such pressure as will be evident to a skilled artisan.

In the embodiment of the present invention illustrated in FIG. 6A, thebeds of solid metal oxide particles and solid metal bromide particlescontained in reactors 240 and 246, respectively, are fluidized and areconnected in the manner described below to provide for continuousoperation of the beds without the need to provide for equipment, such asvalves, to change flow direction to and from each reactor. In accordancewith this embodiment, the effluent which comprises olefins, the highermolecular weight hydrocarbons and hydrobromic acid is withdrawn from thesecond reactor 234 via line 235, cooled to a temperature in the range ofabout 100° C. to about 500° C. in exchanger 236, and introduced into thebottom of reactor 240 which contains a bed 298 of solid metal oxideparticles. The flow of this introduced fluid induces the particles inbed 298 to move upwardly within reactor 240 as the hydrobromic acid isreacted with the metal oxide in the manner as described above withrespect to FIG. 4A. At or near the top of the bed 298, the particleswhich contain substantially solid metal bromide on theattrition-resistant support due to the substantially complete reactionof the solid metal oxide with hydrobromic acid in reactor 240 arewithdrawn via a weir or cyclone or other conventional means of solid/gasseparation, flow by gravity down line 259 and are introduced at or nearthe bottom of the bed 299 of solid metal bromide particles in reactor246. In the embodiment illustrated in FIG. 6A, oxygen, oxygen enrichedair or air 210 is delivered via blower or compressor 213 at a pressurein the range of about ambient to about 10 bar, transported via line 214through heat exchanger 215, wherein the oxygen, oxygen enriched air orair is preheated to a temperature in the range of about 100° C. to about500° C. and introduced into second vessel or reactor 246 below bed 299of a solid phase metal bromide. Oxygen reacts with the metal bromide inthe manner described above with respect to FIG. 4A to produce a dry,substantially HBr free bromine vapor. The flow of this introduced gasinduces the particles in bed 299 to flow upwardly within reactor 246 asoxygen is reacted with the metal bromide. At or near the top of the bed298, the particles which contain substantially solid metal oxide on theattrition-resistant support due to the substantially complete reactionof the solid metal bromide with oxygen in reactor 246 are withdrawn viaa weir or cyclone or other conventional means of solid/gas separation,flow by gravity down line 264 and are introduced at or near the bottomof the bed 298 of solid metal oxide particles in reactor 240. In thismanner, reactors 240 and 246 can be operated continuously withoutchanging the parameters of operation.

In the embodiment illustrated in FIG. 6B, oxygen is utilized as theoxidizing gas and is transported in via line 210 to reactor 246.Accordingly, the embodiment of the process of the present inventionillustrated in FIG. 6A is modified such that the bromine vapor producedfrom reactor 246 is transported via lines 242 and 225 directly to firstreactor 230. Since oxygen is reactive and will not build up in thesystem, the need to condense the bromine vapor to a liquid to removeunreactive components, such as nitrogen, is obviated. Compressor 213 isnot illustrated in FIG. 6B as substantially all commercial sources ofoxygen, such as a commercial air separator unit, will provide oxygen toline 210 at the required pressure. If not, a compressor 213 could beutilized to achieve such pressure as will be evident to a skilledartisan.

In accordance with another embodiment of the process of the presentinvention that is illustrated in FIG. 7, the alkyl bromination and alkylbromide conversion stages are operated in a substantially similar mannerto those corresponding stages described in detail with respect to FIG.4A except as discussed below. Residual air or oxygen and bromine vaporemanating from reactor 246 is transported via line 247, valve 248 andline 242 and valve 300 to heat exchanger or condenser 221 wherein thebromine-containing gas is cooled to a temperature in the range of about30° C. to about 300° C. The bromine-containing vapor is then transportedvia line 242 to vessel or reactor 320 containing a bed 322 of a solidphase metal bromide in a reduced valence state. The metal of the metalbromide in a reduced valence state is selected from copper (Cu), iron(Fe), or molybdenum (Mo). The metal is selected for the impact of itsphysical and thermodynamic properties relative to the desiredtemperature of operation, and also for potential environmental andhealth impacts and cost. Preferably, copper or iron are employed as themetal, with copper being the most preferred. The solid metal bromide ispreferably immobilized on a suitable attrition-resistant support, forexample a synthetic amorphous silica, such as Davicat Grade 57,manufactured by Davison Catalysts of Columbia, Md. More preferably themetal is deposited in oxide form in a range of about 10 to 20 wt % on analumina support with a specific surface area in the range of about 50 to200 m2/g, In reactor 320, bromine vapor is reacted with the solid phasemetal bromide, preferably retained on a suitable attrition-resistantsupport at temperatures below about 300° C. and preferably between about30° C. to about 200° C. in accordance with the following general formulawherein M² represents the metal:2M²Br_(n)+Br₂→2M²Br_(n+1)In this manner, bromine is stored as a second metal bromide, i.e.2M²Br_(n+1), in reactor 320 while the resultant vapor containingresidual air or oxygen is vented from reactor 320 via line 324, valve326 and line 318.

The gas stream containing lower molecular weight alkanes, comprised ofmixture of a feed gas (line 211) and a recycled gas stream, istransported or conveyed via line 262, heat exchanger 352, wherein thegas stream is preheated to a temperature in the range of about 150° C.to about 600° C., valve 304 and line 302 to a second vessel or reactor310 containing a bed 312 of a solid phase metal bromide in an oxidizedvalence state. The metal of the metal bromide in an oxidized valencestate is selected from copper (Cu), iron (Fe), or molybdenum (Mo). Themetal is selected for the impact of its physical and thermodynamicproperties relative to the desired temperature of operation, and alsofor potential environmental and health impacts and cost. Preferably,copper or iron are employed as the metal, with copper being the mostpreferred. The solid metal bromide in an oxidized state is preferablyimmobilized on a suitable attrition-resistant support, for example asynthetic amorphous silica such as Davicat Grade 57, manufactured byDavison Catalysts of Columbia, Md. More preferably the metal isdeposited in an oxide state in a range of 10 to 20 wt % supported on analumina support with a specific surface area of about 50 to 200 m2/g.The temperature of the gas stream is from about 150° C. to about 600°C., and preferably from about 200° C. to about_(—)450° C. In secondreactor 310, the temperature of the gas stream thermally decomposes thesolid phase metal bromide in an oxidized valence state to yieldelemental bromine vapor and a solid metal bromide in a reduced state inaccordance with the following general formula wherein M² represents themetal:2M²Br_(n+1)→2M²Br_(n)+Br₂The resultant bromine vapor is transported with the gas streamcontaining lower molecular weight alkanes via lines 314, 315, valve 317,line 330, heat exchanger 226 prior to being introduced into alkylbromination reactor 230.

Reactors 310 and 320 may be operated in a cyclic fashion. As illustratedin FIG. 7, valve 304 is operated in the open mode to permit the gasstream containing lower molecular weight alkanes to be transported tothe second reactor 310, while valve 317 is operated in the open mode topermit this gas stream with bromine vapor that is generated in reactor310 to be transported to alkyl bromination reactor 230. Likewise, valve306 is operated in the open mode to permit bromine vapor from reactor246 to be transported to reactor 320, while valve 326 is operated in theopen mode to permit residual air or oxygen to be vented from reactor320. Once significant conversion of the reduced metal bromide andoxidized metal bromide in reactors 320 and 310, respectively, to thecorresponding oxidized and reduced states has occurred, these valves areclosed as illustrated in FIG. 8. At this point, bed 322 in reactor 320is a bed of substantially metal bromide in an oxidized state, while bed312 in reactor 310 is substantially metal bromide in a reduced state. Asillustrated in FIG. 8, valves 304, 317, 306 and 326 are closed, and thenvalves 308 and 332 are opened to permit the gas stream containing lowermolecular weight alkanes to be transported or conveyed via lines 262,heat exchanger 352, wherein gas stream is heated to a range of about150° C. to about 600° C., valve 308 and line, 309 to reactor 320 tothermally decompose the solid phase metal bromide in an oxidized valencestate to yield elemental bromine vapor and a solid metal bromide in areduced state. Valve 332 is also opened to permit the resultant brominevapor to be transported with the gas stream containing lower molecularweight alkanes via lines 324 and 330 and heat exchanger 226 prior tobeing introduced into alkyl bromination reactor 230. In addition, valve300 is-opened to permit bromine vapor emanating from reactor 246 to betransported via line 242 through exchanger 221 into reactor 310 whereinthe solid phase metal bromide in a reduced valence state reacts withbromine to effectively store bromine as a metal bromide. In addition,valve 316 is opened to permit the resulting gas, which is substantiallydevoid of bromine to be vented via lines 314 and 318. The reactors areoperated in this manner until significant conversion of the beds ofreduced metal bromide and oxidized metal bromide in reactors 310 and320, respectively, to the corresponding oxidized and reduced states hasoccurred and then the reactors are cycled back to the flow schematicillustrated in FIG. 7 by opening and closing valves as previouslydiscussed.

In the embodiment of the present invention illustrated in FIG. 9, thebeds 312 and 322 contained in reactors 310 and 320, respectively, arefluidized and are connected in the manner described below to provide forcontinuous operation of the beds without the need to provide forequipment, such as valves, to change flow direction to and from eachreactor. In accordance with this embodiment, the bromine-containing gaswithdrawn from the reactor 246 via line 242 is cooled to a temperaturein the range of about 30° C. to about 300° C. in exchangers 370 and 372,and introduced into the bottom of reactor 320 which contains a movingsolid bed 322 in a fluidized state. The flow of this introduced fluidinduces the particles in bed 322 to flow upwardly within reactor 320 asthe bromine vapor is reacted with the reduced metal bromide entering thebottom of bed 322 in the manner as described above with respect to FIG.7. At or near the top of the bed 322, the particles which containsubstantially oxidized metal bromide on the attrition-resistant supportdue to the substantially complete reaction of the reduced metal bromidewith bromine vapor in reactor 320 are withdrawn via a weir, cyclone orother conventional means of solid/gas separation, flow by gravity downline 359 and are introduced at or near the bottom of the bed 312 inreactor 310. In the embodiment illustrated in FIG. 9, the gas streamcontaining lower molecular weight alkanes, comprised of mixture of afeed gas (line 211) and a recycled gas stream, is transported orconveyed via line 262 and heat exchanger 352 wherein the gas stream isheated to a range of about 150° C. to about 600° C. and introduced intoreactor 310. The heated gas stream thermally decomposes the solid phasemetal bromide in an oxidized valence state present entering at or nearthe bottom of bed 312 to yield elemental bromine vapor and a solid metalbromide in a reduced state. The flow of this introduced gas induces theparticles in bed 312 to flow upwardly within reactor 310 as the oxidizedmetal bromide is thermally decomposed. At or near the top of the bed312, the particles which contain substantially reduced solid metalbromide on the attrition-resistant support due to the substantiallycomplete thermal decomposition in reactor 310 are withdrawn via a weiror cyclone or other conventional means of gas/solid separation and flowby gravity down line 364 and introduced at or near the bottom of the bed322 of particles in reactor 310. In this manner, reactors 310 and 320can be operated continuously with changing the parameters of operation.

The process of the present invention is less expensive than conventionalprocess since it operates at low pressures in the range of about 1 barto about 30 bar and at relatively low temperatures in the range of about20° C. to about 600° C. for the gas phase, and preferably about 20° C.to about 180° C. for the liquid phase. These operating conditions permitthe use of less expensive equipment of relatively simple design that areconstructed from readily available metal alloys or glass-lined equipmentfor the gas phase and polymer-lined or glass-lined vessels, piping andpumps for the liquid phase. The process of the present invention is alsomore efficient because less energy is required for operation and theproduction of excessive carbon dioxide as an unwanted byproduct isminimized. The process is capable of directly producing a mixedhydrocarbon product containing various molecular-weight components inthe liquefied petroleum gas (LPG), olefin and motor gasoline fuels rangethat have substantial aromatic content thereby significantly increasingthe octane value of the gasoline-range fuel components.

While the foregoing preferred embodiments of the invention have beendescribed and shown, it is understood that the alternatives andmodifications, such as those suggested and others, may be made theretoand fall within the scope of the invention.

1. A process for converting gaseous alkanes to olefins and liquidhydrocarbons comprising: reacting a gaseous feed having lower molecularweight alkanes with bromine vapor to form alkyl bromides and hydrobromicacid; and reacting said alkyl bromides and hydrobromic acid in thepresence of a synthetic crystalline alumino-silicate catalyst and at atemperature sufficient to form olefins and hydrobromic acid vapor. 2.The process of claim 1 wherein said bromine vapor is substantially dry,thereby avoiding the formation of significant carbon dioxide along withsaid alkyl bromides.
 3. The process of claim 1 wherein said gaseous feedis natural gas.
 4. The process of claim 3 wherein said natural gas istreated to remove substantially all of the carbon dioxide and sulfurcompounds therefrom prior to reacting with said bromine vapor.
 5. Theprocess of claim 1 wherein said temperature is from about 250° C. toabout 500° C.
 6. The process of claim 5 wherein said temperature is fromabout 300° C. to about 450° C.
 7. The process of claim 1 wherein saidcrystalline alumino-silicate catalyst is a zeolite catalyst.
 8. Theprocess of claim 7 wherein said zeolite catalyst is an X or Y typezeolite catalyst.
 9. The process of claim 8 wherein said zeolitecatalyst is a 10 X zeolite catalyst.
 10. The process of claim 1 furthercomprising: removing said hydrobromic acid vapor from said olefins byneutralization reaction with an aqueous solution containing reactionproducts obtained by oxidizing an aqueous solution containing a metalbromide salt, the metal of said metal bromide salt being selected fromCu, Zn, Fe, Co, Ni, Mn, Ca or Mg bromide.
 11. The process of claim 1wherein said bromine vapor is produced by oxidizing an aqueous metalbromide salt solution, the metal of said metal bromide salt beingselected from Cu, Zn, Fe, Co, Ni, Mn, Ca, or Mg.
 12. A process forconverting gaseous lower molecular weight alkanes to olefins comprising:reacting a gaseous feed containing lower molecular weight alkanes withbromine vapor to form alkyl bromides and hydrobromic acid; reacting saidalkyl bromides and hydrobromic acid in the presence of a syntheticcrystalline alumino-silicate catalyst to form olefins and hydrobromicacid; and converting said hydrobromic acid to bromine.
 13. The processof claim 12 further comprising: dehydrating said olefins.
 14. Theprocess of claim 12 further comprising: recycling said bromine that isconverted from said hydrobromic acid to said step of reacting with saidgaseous feed, said bromine being recycled as a vapor.
 15. A process forconverting gaseous alkanes to olefins comprising: reacting a gaseousfeed having lower molecular weight alkanes with bromine vapor to formalkyl bromides and hydrobromic acid; reacting said alkyl bromides andhydrobromic acid in the presence of a synthetic crystallinealumino-silicate catalyst and at a temperature sufficient to formolefins and hydrobromic acid vapor; and removing said hydrobromic acidvapor from said olefins by reacting said hydrobromic acid vapor with ametal oxide to form a metal bromide and steam.
 16. The process of claim15 wherein said the metal of said metal oxide is magnesium, calcium,vanadium, chromium, manganese, iron, cobalt, nickel, copper, zinc ortin.
 17. The process of claim 15 wherein said metal oxide is supportedon a solid carrier.
 18. The process of claim 15 wherein said metal oxideis contained in a bed in a reactor.
 19. The process of claim 15 furthercomprising: reacting said metal bromide with an oxygen containing gas toobtain a metal oxide and said bromine vapor.
 20. The process of claim 19wherein said bromine vapor is used in the step of reacting said gaseousfeed having lower molecular weight alkanes.